Process for improving the selectivity of an EO catalyst

ABSTRACT

The present invention relates to a process for improving the overall selectivity of an EO process for converting ethylene to ethylene oxide utilizing a highly selective EO silver catalyst containing a rhenium promoter wherein following normal operation a hard strip of the chloride on the surface of the catalyst is conducted in order to remove a portion of the chlorides on the surface of the catalyst. Following the hard strip, the catalyst is optionally re-optimized. Surprisingly, it has been found that the selectivity of the catalyst following the hard strip may be substantially higher than the selectivity prior to the hard strip.

This application claims the benefit of U.S. Provisional Application61/420,844 filed Dec. 8, 2010, which is herein incorporated byreference.

FIELD OF THE INVENTION

The invention relates to a process for the operation of an ethyleneepoxidation process which employs a silver-based highly selectiveepoxidation catalyst. The invention also relates to a process for theproduction of ethylene oxide, a 1,2-diol, a 1,2-diol ether, a1,2-carbonate, or an alkanolamine.

BACKGROUND OF THE INVENTION

In olefin epoxidation an olefin is reacted with oxygen to form an olefinepoxide, using a catalyst comprising a silver component, usually withone or more further elements deposited therewith on a support. Theolefin oxide may be reacted with water, an alcohol, carbon dioxide or anamine to form a 1,2-diol, a 1,2-diol ether, 1,2 carbonate or analkanolamine. Thus, 1,2-diols, 1,2-diol ethers and alkanolamines may beproduced in a multi-step process comprising olefin epoxidation andconverting the formed olefin oxide with water, an alcohol or an amine.

The performance of the epoxidation process may be assessed on the basisof the selectivity, the catalyst's activity, and stability of operation.The selectivity is the molar fraction of the converted olefin yieldingthe desired olefin oxide. Modern silver-based epoxidation catalysts arehighly selective towards olefin oxide production. When using the moderncatalysts in the epoxidation of ethylene the selectivity towardsethylene oxide can reach values above 85 mole-%. An example of suchhighly selective catalysts is a catalyst comprising silver and a rheniumpromoter, for example U.S. Pat. Nos. 4,761,394 and 4,766,105.

For decades much research has been devoted to improving the activity,the selectivity, and the lifetime of the catalysts, and to find processconditions which enable full exploitation of the catalyst performance. Areaction modifier, for example an organic halide, may be added to thefeed in an epoxidation process for increasing the selectivity of ahighly selective catalyst (see for example EP-A-352850, U.S. Pat. Nos.4,761,394 and 4,766,105, which are herein incorporated by reference).The reaction modifier suppresses the undesirable oxidation of olefin orolefin oxide to carbon dioxide and water, relative to the desiredformation of olefin oxide. EP-A-352850 teaches that there is an optimumin the selectivity as a function of the quantity of organic halide inthe feed, at a constant oxygen conversion level and given set ofreaction conditions.

Many process improvements are known that can improve selectivity. Forexample, it is well known that low CO₂ levels are useful in improvingthe selectivity of high selectivity catalysts. See, e.g., U.S. Pat. Nos.7,235,677; 7,193,094; US Pub. Pat. App. 2007/0129557; WO 2004/078736; WO2004/078737; and EP 2,155,708. These patents also disclose that waterconcentration in the reactor feed should be maintained at a level of atmost 0.35 mole percent, preferably less than 0.2 mole percent. Otherpatents disclose control of the chloride moderator to maintain goodactivity. See, e.g., U.S. Pat. No. 7,657,331; EP 1,458,698; and U.S.Pub. Pat. App. 2009/0069583. Still further, there are many other patentsdealing with EO process operation and means to improve the performanceof the catalyst in the process. See, e.g., U.S. Pat. Nos. 7,485,597,7,102,022, 6,717,001, 7,348,444, and U.S. Pub. Pat. App. 2009/0234144.

All catalysts must first be started up in a manner to establish a goodselectivity operation. U.S. Pat. No. 7,102,022 relates to the start-upof an epoxidation process wherein a highly selective catalyst isemployed. In this patent there is disclosed an improved start-upprocedure wherein the highly selective catalyst is subjected to a heattreatment wherein the catalyst is contacted with a feed comprisingoxygen at a temperature above the normal operating temperature of thehighly selective catalyst (i.e., above 260° C.). U.S. Pub. Pat. App.2004/0049061 relates to a method of improving the selectivity of ahighly selective catalyst having a low silver density. In this document,there is disclosed a method wherein the highly selective catalyst issubjected to a heat treatment which comprises contacting the catalystwith a feed comprising oxygen at a temperature above the normaloperating temperature of the highly selective catalyst (i.e., above 250°C.). U.S. Pat. No. 4,874,879 relates to the start-up of an epoxidationprocess employing a highly selective catalyst wherein the highlyselective catalyst is first contacted with a feed containing an organicchloride moderator and ethylene, and optionally a ballast gas, at atemperature below the normal operating temperature of the catalyst.EP-B1-1532125 relates to an improved start-up procedure wherein thehighly selective catalyst is first subjected to a pre-soak phase in thepresence of a feed containing an organic halide and is then subjected toa stripping phase in the presence of a feed which is free of the organichalide or may comprise the organic halide in a low quantity. Thestripping phase is taught to continue for a period of more than 16 hoursup to 200 hours. U.S. Pat. App. No. 2009/0281339 relates to the start-upwhere the organic chloride in the feed is adjusted to a value sufficientto produce EO at a substantially optimum selectivity.

At the end of the start-up period, the chloride level is typicallyadjusted to find the chloride level which gives the maximum selectivityat the desired EO production rate. The plant then sets the chloridelevel equal to this so-called “chloride optimum” and begins normaloperation of the catalyst, which continues until it is discharged fromthe reactor. During normal operation of the catalyst, several routinethings may happen:

-   -   The catalyst will deactivate. In order to maintain a constant        production rate, the reaction temperature will be increased as        the catalyst deactivates.    -   The production rate may change, due to feedstock availability,        production demands, or economics. To increase the production        rate, the reaction temperature will be increased; to decrease        the production rate, the reaction temperature will be decreased.    -   The feed composition may change. Generally, CO₂ levels will        increase over the life of the catalyst as selectivity drops.        Also, ethylene and oxygen levels may be changed due to feedstock        issues or to lower temperature near end-of-cycle.    -   Feed impurities (such as ethane or propane) may fluctuate.    -   There may be an upset in operation due to such events as        equipment failure or unplanned operation changes or deviations        from normal operation.

It is well-known (see, e.g., U.S. Pat. No. 7,193,094 and EP 1,458,698)that changes in reaction temperature or hydrocarbon concentration willchange the chloride optimum. For example, as the reaction temperatureincreases or as hydrocarbon levels increase, the chloride level willalso need to be increased in order to maintain operation at the maximumselectivity. During routine plant operation, the chloride level isadjusted in one of two methods:

-   -   1. The plant utilizes some proprietary mathematical formula        which relates chloride level to temperature, composition, etc.        This formula is computed periodically and if the chloride level        is found to be significantly different than the optimal level        (as determined by the formula), then the chloride level is        adjusted so that it equals the optimal level.    -   2. More frequently, the plant routinely checks whether the        chloride level is still optimized. This may happen at some fixed        frequency or following certain changes in operating conditions,        as determined by the plant. Typically, the chloride level is        increased or decreased slightly and the plant observes whether        the selectivity changed. If it did not change, then they were        probably operating at the selectivity maximum, so the chloride        level is reset to its original value. If the selectivity did        change, then the chloride level is changed in small steps until        a selectivity maximum is found, and then the plant continues        operation at this new chloride optimum.

Notwithstanding the improvements already achieved, there is a desire tofurther improve the performance of the silver-containing catalysts inthe production of an olefin oxide, a 1,2-diol, a 1,2-diol ether,1,2-carbonate or an alkanolamine.

SUMMARY OF THE INVENTION

The present invention shows that routine operation of a high-selectivitycatalyst in the typical fashion described above is not truly optimal. Itis believed that the processes that occur during start-up, namelyexposure of the catalyst to high temperatures in the presence of lowchloride levels, are necessary to distribute the catalyst dopantsoptimally. Incremental changes to the chloride level over the course ofa catalyst run do not usually provide the proper conditions to maintainthe optimal surface dopant configuration.

Significant increases in the selectivity of a catalyst may be obtainedthrough conducting a “hard strip” according to the present invention, inwhich the feed chloride level is significantly reduced or eliminated fora period of time. After the hard strip, the chloride level is firstreturned to around the prior chloride level (i.e., the chloride level inuse immediately prior to the hard strip). In one option the catalyst isre-optimized at that point in time. It was found that the claimed hardstrip procedure may increase the catalyst selectivity by severalselectivity points, with the greatest benefit being for catalysts thathave lost significant selectivity over the course of the run due toaging. In such cases, selectivity gains as high as 3% were observedfollowing a hard strip.

In accordance with this invention, the operation of an epoxidationprocess using a highly selective catalyst can be improved by utilizingthe process steps according to the present invention. In particular, thepresent invention comprises a process for improving the selectivity ofan ethylene epoxidation process employed in a reactor comprising acatalyst bed having a multitude of reactor tubes filled with a highselectivity epoxidation catalyst, said process comprising:

-   -   (a) contacting the catalyst bed with a feed comprising ethylene,        oxygen, and an organic chloride moderator for at least a period        of time T₁ to produce ethylene oxide;    -   (b) subsequently subjecting the high selectivity epoxidation        catalyst to a hard strip over a period of time T₂, which        comprises:        -   (i) significantly reducing the organic chloride added to the            feed; and        -   (ii) treating the catalyst in order to strip a portion of            the chlorides from the surface of the catalyst by increasing            the reactor temperature; and    -   (c) following the hard strip, increasing the quantity of organic        chloride added to the feed to a level of between 80 and 120% of        the prior concentration and reducing the reactor temperature to        within 5° C. of the prior reactor temperature level.

Step (a) of the present invention comprises the normal operation of theprocess to produce ethylene oxide according to the design parameters ofthe plant and process. According to the present invention, it ispossible to improve operation and selectivity by inclusion of anintermediate stripping step (b), where the organic chloride feed issignificantly reduced or stopped and the reactor temperature increasedfor a period of time until the chloride level on the catalyst surface issignificantly reduced.

In an optional step following step (c) the various operating conditionsmay be re-optimized to achieve a new optimum selectivity prior toresuming normal operation. Normal operation would comprise subsequentlycontacting the catalyst bed with feed comprising ethylene, oxygen, andan organic chloride moderator.

The organic chloride for use in the present process is typically one ormore chloro-hydrocarbons. Preferably, the organic chloride is selectedfrom the group of methyl chloride, ethyl chloride, ethylene dichloride,vinyl chloride or a mixture thereof. The most preferred reactionmodifiers are ethyl chloride, vinyl chloride and ethylene dichloride.

The selectivity (to ethylene oxide) indicates the molar amount ofethylene oxide in the reaction product compared with the total molaramount of ethylene converted. By high-selectivity is meant a catalystwith selectivity greater than 80 molar-%, preferably greater than 85.7molar-%. One such catalyst is a rhenium containing catalyst, such asthat disclosed in U.S. Pat. No. 4,766,105 or US 2009/0281345A1.

T₁ refers to the initial period of time following start-up and isdefined as the time required to produce at least 0.1 kilotons (10⁵ kg)of ethylene oxide per cubic meter of catalyst. For the typical range ofcommercial workrates from 100 to 300 kg/m³/hour, this translates toabout 14 to about 42 days. T₂ refers to the time period for the chloridestrip and should be as short as possible in order to start producing EOat the same or better commercial rates as before the chloride strip.This T₂ time period is typically about 2 to 72 hours, more preferablyabout 4 to about 24 hours.

It appears that the long used process of initial optimization of thecatalyst does not produce a catalyst that will remain optimum over thelife of the catalyst. Rather, it appears that the distribution ofchlorides on the catalyst cannot reach a true optimum until there is astripping of the chlorides, followed by a reintroduction of thechlorides. This constitutes a previously unknown operational method ofsignificantly improving EO production profitability. Stripping thecatalyst is common practice during start-up, but subsequent strippingdoes not happen in normal operation. This is because during a chloridestrip, selectivity and EO production both drop, leading to reducedproduction by the EO plant, and plant operators are unwilling to losethat production of EO. However, as shown in the present invention, theoccasional “hard strip” of chlorides done according to the inventionreturns at least 0.5% or better selectivity improvement for an extendedperiod of operation. Therefore, the implementation of occasional hardstrips may constitute the preferred mode of operation of the EO catalystin the future.

DETAILED DESCRIPTION OF THE INVENTION

Although the present epoxidation process may be carried out in manyways, it is preferred to carry it out as a gas phase process, i.e. aprocess in which the feed is contacted in the gas phase with thecatalyst which is present as a solid material, typically in a packedbed. Generally the process is carried out as a continuous process. Thereactor is typically equipped with heat exchange facilities to heat orcool the catalyst. As used herein, the feed is considered to be thecomposition which is contacted with the catalyst. As used herein, thecatalyst temperature or the temperature of the catalyst bed is deemed tobe the temperature approximately half-way through the catalyst bed.

When new catalysts as well as aged catalysts which, due to a plantshut-down, have been subjected to a prolonged shut-in period areutilized in the epoxidation process, it may be useful in some instancesto pre-treat these catalysts prior to carrying out the start-up processby passing a sweeping gas over the catalyst at an elevated temperature.The sweeping gas is typically an inert gas, for example nitrogen orargon, or mixtures comprising nitrogen and/or argon. The elevatedtemperature converts a significant portion of organic nitrogen compoundswhich may have been used in the manufacture of the catalyst to nitrogencontaining gases which are swept up in the gas stream and removed fromthe catalyst. In addition, any moisture may be removed from thecatalyst. Typically, when the catalyst is loaded into the reactor, byutilizing the coolant heater, the temperature of the catalyst is broughtup to 200 to 250° C., preferably from 210 to 230° C., and the gas flowis passed over the catalyst. Further details on this pre-treatment maybe found in U.S. Pat. No. 4,874,879, which is incorporated herein byreference.

The catalyst is subjected to a start-up process which involves aninitial step of contacting the catalyst with a feed comprising ethylene,oxygen, and an organic chloride. For the sake of clarity only, this stepof the process will be indicated hereinafter by the term “initialstart-up phase”. During the initial start-up phase, the catalyst is ableto produce ethylene oxide at or near the selectivity experienced afterthe catalyst has “lined-out” under normal initial operating conditionsafter the start-up process. In particular, during the initial start-upphase, the selectivity may be within 3 mole-%, more in particular within2 mole-%, most in particular within 1 mole-% of the optimum selectivityperformance under normal initial operating conditions. Suitably, theselectivity may reach and be maintained at more than 86.5 mole-%, inparticular at least 87 mole-%, more in particular at least 87.5 mole-%during the initial start-up phase. Since the selectivity of the catalystquickly increases, there is advantageously additional production ofethylene oxide.

In the initial start-up phase, the catalyst is contacted with organicchloride for a period of time until an increase of at least 1×10⁻⁵mole-% of vinyl chloride (calculated as the moles of vinyl chloriderelative to the total gas mixture) is detected in the reactor outlet orthe recycle gas loop. Without wishing to be bound by theory, when usingorganic chlorides other than vinyl chloride, it is believed that thevinyl chloride detected in the outlet or recycle loop is generated bythe reaction of surface adsorbed chloride on the silver present in thecatalyst with a C₂ hydrocarbon present in the feed. Preferably, thecatalyst is contacted with organic chloride for a period of time untilan increase of at least 2×10⁻⁵ mole-% of vinyl chloride, in particularat most 1×10⁻⁴ mole-% (calculated as the moles of vinyl chloriderelative to the total gas mixture) is detected in the reactor outlet orthe recycle gas loop. The quantity of organic chloride contacted withthe catalyst during the initial start-up phase may be in the range offrom 1 to 12 millimolar (mmolar) equivalent of chloride per kilogram ofcatalyst. The mmolar equivalent of chloride is determined by multiplyingthe mmoles of the organic chloride by the number of chloride atomspresent in the organic chloride molecule, for example 1 mmole ofethylene dichloride provides 2 mmolar equivalent of chloride. Theorganic chloride may be fed to the catalyst bed for a period of timeranging from 1 to 15 hours, preferably 2 to 10 hours, more preferablyfrom 2.5 to 8 hours. Suitably, the quantity of the organic chloridecontacted with the catalyst may be at most 6 mmolar equivalent/kgcatalyst, in particular at most 5.5 mmolar equivalent/kg catalyst, morein particular at most 5 mmolar equivalent/kg catalyst. The quantity ofthe organic chloride in the feed during the initial start-up phase maybe at least 1.5×10⁻⁴ mole-%, in particular at least 2×10 mole-%,calculated as moles of chloride, relative to the total feed. Thequantity of the organic chloride during the initial start-up phase maybe at most 0.1 mole-%, preferably at most 0.01 mole-%, calculated asmoles of chloride, relative to the total feed. Preferably, the initialstart-up feed may comprise the organic chloride in a quantity above theoptimum quantity used during the initial period of normal ethylene oxideproduction.

The feed during the initial start-up phase may also contain additionalreaction modifiers which are not organic halides, such as nitrate- ornitrite-forming compounds, as described herein.

The feed during the initial start-up phase also contains ethylene.Ethylene may be present in the initial start-up feed in a quantity of atleast 10 mole-%, preferably at least 15 mole-%, more preferably at least20 mole-%, relative to the total feed. Ethylene may be present in theinitial start-up feed in a quantity of at most 50 mole-%, preferably atmost 45 mole-%, more preferably at most 40 mole-%, relative to the totalfeed. Preferably, ethylene may be present in the initial start-up feedin the same or substantially the same quantity as utilized during normalethylene oxide production. This provides an additional advantage in thatethylene concentration does not have to be adjusted between the initialstart-up phase and normal ethylene oxide production post start-up makingthe process more efficient.

The feed during the initial start-up phase also contains oxygen. Theoxygen may be present in the initial start-up feed in a quantity of atleast 1 mole-%, preferably at least 2 mole-%, more preferably at least2.5 mole-%, relative to the total feed. The oxygen may be present in theinitial start-up feed in a quantity of at most 15 mole-%, preferably atmost 10 mole-%, more preferably at most 5 mole-%, relative to the totalfeed. It may be advantageous to apply a lower oxygen quantity in theinitial start-up feed, compared with the feed composition in laterstages of the process during normal ethylene oxide production since alower oxygen quantity in the feed will reduce the oxygen conversionlevel so that, advantageously, hot spots in the catalyst are betteravoided and the process will be more easily controllable.

The feed during the initial start-up phase may also contain carbondioxide. The carbon dioxide may be present in the initial start-up feedin a quantity of at most 10 mole-%, preferably at most 5 mole-%,relative to the total feed. In an embodiment, the initial start-up phasealso contains less than 2 mole-%, preferably less than 1.5 mole percent,more preferably less than 1.2 mole percent, most preferably less than 1mole percent, in particular at most 0.75 mole percent carbon dioxide,relative to the total feed. In the normal practice of the presentinvention, the quantity of carbon dioxide present in the reactor feed isat least 0.1 mole percent, or at least 0.2 mole percent, or at least 0.3mole percent, relative to the total feed. Suitably, the carbon dioxidemay be present in the initial start-up feed in the same or substantiallythe same quantity as utilized during normal ethylene oxide production.The balance of the feed during the initial start-up phase may alsocontain an inert and/or saturated hydrocarbon.

During the initial start-up phase, the catalyst temperature preferablymay be at substantially the same temperature as the normal initialcatalyst operating temperature after the epoxidation process has“lined-out” under normal operating conditions after the start-upprocess. The term “substantially the same temperature” as used herein ismeant to include catalyst temperatures within ±5° C. of the normalinitial catalyst operating temperature after the epoxidation process has“lined-out” under normal operating conditions after the start-upprocess. Preferably, the catalyst temperature is less than 250° C.,preferably at most 245° C. The catalyst temperature may be at least 200°C., preferably at least 220° C., more preferably at least 230° C. Thereactor inlet pressure may be at most 4000 kPa absolute, preferably atmost 3500 kPa absolute, more preferably at most 2500 kPa absolute. Thereactor inlet pressure is at least 500 kPa absolute. The Gas HourlySpace Velocity or “GHSV”, defined hereinafter, may be in the range offrom 500 to 10000 N1/(1. h).

During the initial start-up phase, the catalyst may first be contactedwith a feed comprising ethylene and optionally a saturated hydrocarbon,in particular ethane and optionally methane. The organic chloride maythen be added to the feed. The oxygen may be added to the feedsimultaneously with or shortly after the first addition of the organicchloride to the feed. Within a few minutes of the addition of oxygen,the epoxidation reaction can initiate. Carbon dioxide and additionalfeed components may be added at any time, preferably simultaneously withor shortly after the first addition of oxygen to the initial start-upfeed. As discussed above, during the initial start-up phase, thecatalyst is able to produce ethylene oxide at or near the selectivityexperienced after the catalyst has “lined-out” under normal initialoperating conditions after the start-up process. During the initialstart-up phase, the catalyst is operated under conditions such thatethylene oxide is produced at a level that is from 45 to 100% of thetargeted production level during normal ethylene oxide production, inparticular from 50 to 70%, same basis.

The present epoxidation process may be air-based or oxygen-based, see“Kirk-Othmer Encyclopedia of Chemical Technology”, 3^(rd) edition,Volume 9, 1980, pp. 445-447. In the air-based process, air or airenriched with oxygen is employed as the source of the oxidizing agentwhile in the oxygen-based processes, high-purity (at least 95 mole-%) orvery high purity (at least 99.5 mole-%) oxygen is employed as the sourceof the oxidizing agent. Reference may be made to U.S. Pat. No.6,040,467, incorporated by reference, for further description ofoxygen-based processes. Presently most epoxidation plants areoxygen-based and this is a preferred embodiment of the presentinvention.

In addition to ethylene, oxygen and the organic chloride, the productionfeed during the normal epoxidation process may contain one or moreoptional components, such as nitrogen-containing reaction modifiers,carbon dioxide, inert gases and saturated hydrocarbons.

Nitrogen oxides, organic nitro compounds such as nitromethane,nitroethane, and nitropropane, hydrazine, hydroxylamine or ammonia maybe employed as reaction modifiers in the epoxidation process. It isfrequently considered that under the operating conditions of ethyleneepoxidation the nitrogen containing reaction modifiers are precursors ofnitrates or nitrites, i.e. they are so-called nitrate- ornitrite-forming compounds. Reference may be made to EP-A-3642 and U.S.Pat. No. 4,822,900, which are incorporated herein by reference, forfurther description of nitrogen-containing reaction modifiers.

Suitable nitrogen oxides are of the general formula NO_(x) wherein x isin the range of from 1 to 2.5, and include for example NO, N₂O₃, N₂O₄,and N₂O₅. Suitable organic nitrogen compounds are nitro compounds,nitroso compounds, amines, nitrates and nitrites, for examplenitromethane, 1-nitropropane or 2-nitropropane.

Carbon dioxide is a by-product in the epoxidation process. However,carbon dioxide generally has an adverse effect on the catalyst activity,and high concentrations of carbon dioxide are therefore typicallyavoided. A typical epoxidation reactor feed during the normalepoxidation process may contain a quantity of carbon dioxide in the feedof at most 10 mole-%, relative to the total feed, preferably at most 5mole-%, relative to the total feed. A quantity of carbon dioxide of lessthan 3 mole-%, preferably less than 2 mole-%, more preferably less than1 mole-%, relative to the total feed, may be employed. Under commercialoperations, a quantity of carbon dioxide of at least 0.1 mole-%, inparticular at least 0.2 mole-%, relative to the total feed, may bepresent in the feed.

The inert gas may be, for example, nitrogen or argon, or a mixturethereof. Suitable saturated hydrocarbons are propane and cyclopropane,and in particular methane and ethane. Saturated hydrocarbons may beadded to the feed in order to increase the oxygen flammability limit.

In the normal ethylene oxide production phase, the invention may bepracticed by using methods known in the art of epoxidation processes.For further details of such epoxidation methods reference may be made,for example, to U.S. Pat. Nos. 4,761,394, 4,766,105, 6,372,925,4,874,879, and 5,155,242, which are incorporated herein by reference.

In normal ethylene oxide production phase, the process may be carriedout using reaction temperatures selected from a wide range. Preferablythe reaction temperature is in the range of from 150 to 325° C., morepreferably in the range of from 180 to 300° C.

In the normal ethylene oxide production phase, the concentration of thecomponents in the feed may be selected within wide ranges, as describedhereinafter.

The quantity of ethylene present in the production feed may be selectedwithin a wide range. The quantity of ethylene present in the feed willbe at most 80 mole-%, relative to the total feed. Preferably, it will bein the range of from 0.5 to 70 mole-%, in particular from 1 to 60mole-%, on the same basis. Preferably, the quantity of ethylene in theproduction feed is substantially the same as used in the start-upprocess. If desired, the ethylene concentration may be increased duringthe lifetime of the catalyst, by which the selectivity may be improvedin an operating phase wherein the catalyst has aged, see U.S. Pat. No.6,372,925 which methods are incorporated herein by reference.

Often present in the feed will be saturated hydrocarbons such as ethaneand methane. These saturated hydrocarbons are also termed“co-moderators”, since they have an impact on the effect of the chloride“moderators”, in that they are effective at removing or “stripping”adsorbed chloride from the surface of the catalyst. The level of ethaneis typically 0.05 to 1.5 mole-% of the feed, more typically 0.05 to 0.5mol-% of the feed, and will depend upon the particular feed stream tothe reactor. Such level is monitored, but is not usually controlledduring normal operation.

The quantity of oxygen present in the production feed may be selectedwithin a wide range. However, in practice, oxygen is generally appliedin a quantity which avoids the flammable regime. The quantity of oxygenapplied will be within the range of from 4 to 15 mole-%, more typicallyfrom 5 to 12 mole-% of the total feed.

In order to remain outside the flammable regime, the quantity of oxygenpresent in the feed may be lowered as the quantity of ethylene isincreased. The actual safe operating ranges depend, along with the feedcomposition, also on the reaction conditions such as the reactiontemperature and the pressure.

The organic chlorides are generally effective as a reaction modifierwhen used in small quantities in the production feed, for example up to0.1 mole-%, calculated as moles of chloride, relative to the totalproduction feed, for example from 0.01×10⁴ to 0.01 mole-%, calculated asmoles of chloride, relative to the total production feed. In particular,it is preferred that the organic chloride may be present in the feed ina quantity of from 1×10⁴ to 50×10⁴ mole-%, in particular from 1.5×10⁴ to25×10⁴ mole-%, more in particular from 1.75×10⁴ to 20×10⁴ mole-%,calculated as moles of chloride, relative to the total production feed.When nitrogen containing reaction modifiers are applied, they may bepresent in low quantities in the feed, for example up to 0.1 mole-%,calculated as moles of nitrogen, relative to the total production feed,for example from 0.01×10⁴ to 0.01 mole-%, calculated as moles ofnitrogen, relative to the total production feed. In particular, it ispreferred that the nitrogen containing reaction modifier may be presentin the feed in a quantity of from 0.05×10⁴ to 50×10⁴ mole-%, inparticular from 0.2×10⁴ to 30×10⁴ mole-%, more in particular from0.5×10⁴ to 10×10⁴ mole-%, calculated as moles of nitrogen, relative tothe total production feed.

Inert gases, for example nitrogen or argon, may be present in theproduction feed in a quantity of 0.5 to 90 mole-%, relative to the totalfeed. In an air based process, inert gas may be present in theproduction feed in a quantity of from 30 to 90 mole-%, typically from 40to 80 mole-%. In an oxygen-based process, inert gas may be present inthe production feed in a quantity of from 0.5 to 30 mole-%, typicallyfrom 1 to 15 mole-%. If saturated hydrocarbons are present, they may bepresent in a quantity of up to 80 mole-%, relative to the totalproduction feed, in particular up to 75 mole-%, same basis. Frequentlythey are present in a quantity of at least 30 mole-%, more frequently atleast 40 mole-%, same basis.

In the normal ethylene oxide production phase, the epoxidation processis preferably carried out at a reactor inlet pressure in the range offrom 1000 to 3500 kPa. “GHSV” or Gas Hourly Space Velocity is the unitvolume of gas at normal temperature and pressure (0° C., 1 atm, i.e.101.3 kPa) passing over one unit volume of packed catalyst per hour.Preferably, when the epoxidation process is a gas phase processinvolving a packed catalyst bed, the GHSV is in the range of from 1500to 10000 N1/(1. h). Preferably, the process is carried out at a workrate in the range of from 0.5 to 10 kmole ethylene oxide produced per m³of catalyst per hour, in particular 0.7 to 8 kmole ethylene oxideproduced per m³ of catalyst per hour, for example 5 kmole ethylene oxideproduced per m³ of catalyst per hour. As used herein, the work rate isthe amount of ethylene oxide produced per unit volume of catalyst perhour and the selectivity is the molar quantity of ethylene oxide formedrelative to the molar quantity of ethylene converted.

The key to the present invention is to initiate a “hard strip” of thechlorides following the initial start-up of the process and after apartial production of the planned EO run, but prior to the planned shutdown and catalyst replacement. In the past, EO plant operators would notconsider a “hard strip”, but would rather alter operations by increasingreactor temperature and/or changing the relative amount of the chloridemoderator in order to improve activity and/or selectivity as thecatalyst degraded over time. This would continue until it becamenecessary to shut down the plant in order to remove the old degradedcatalyst and replace it with a new catalyst. Now it has been discoveredthat there is an economic incentive to introduce a hard strip of thechlorides during the normal operation. Stripping of the chlorides isaccomplished by first stopping or significantly reducing theintroduction of fresh chloride moderator and increasing reactortemperature. The normal level of chloride addition is a function ofcatalyst, temperature, gas flowrate and catalyst volume. Significantlyreducing the amount means to reduce the introduction of fresh chloridemoderator into the feed stream to the reactor by at least 50%, morepreferably at least 75% and most preferably by eliminating all the freshaddition of chloride (100% reduction). In order to strip a portion ofthe chlorides from the surface of the catalyst it will typically also benecessary to increase the reactor temperature by about 1 to about 30°C., preferably by about 2 to about 15° C.

Following the chloride strip, the chloride is then reintroduced over arelatively short time period until the level reaches approximately thesame as prior to the hard strip. As for the time period for the hardstrip, this should be as short as possible to effect the desired resultin order to start producing EO at the same or better commercial rates asbefore the hard strip. This time period is typically about 2 to 72hours, more preferably about 4 to about 24 hours.

Following the hard strip, the process may be continued under normalplant conditions. In an optional step the process may be re-optimizedfollowing the hard strip. This may be done by varying the chloridemoderator up and down to determine at what level the selectivity is atsubstantially an optimum. This optimization step is fully described inU.S. Pat. No. 7,193,094 and US Published Application No. 2009/0281339,which disclosures are hereby incorporated by reference.

Following the first “hard strip”, the hard strip may be repeated laterin the run as needed (e.g., if selectivity drops below expectations).

The epoxidation catalyst is a supported catalyst. The carrier may beselected from a wide range of materials. Such carrier materials may benatural or artificial inorganic materials and they include siliconcarbide, clays, pumice, zeolites, charcoal, and alkaline earth metalcarbonates, such as calcium carbonate. Preferred are refractory carriermaterials, such as alumina, magnesia, zirconia, silica, and mixturesthereof. The most preferred carrier material is a-alumina.

The surface area of the carrier may suitably be at least 0.1 m²/g,preferably at least 0.3 m²/g, more preferably at least 0.5 m²/g, and inparticular at least 0.6 m²/g, relative to the weight of the carrier; andthe surface area may suitably be at most 20 m²/g, preferably at most 10m²/g, more preferably at most 6 m²/g, and in particular at most 4 m²/g,relative to the weight of the carrier. “Surface area” as used herein isunderstood to relate to the surface area as determined by the B.E.T.(Brunauer, Emmett and Teller) method as described in Journal of theAmerican Chemical Society 60 (1938) pp. 309-316. High surface areacarriers, in particular when they are alpha alumina carriers optionallycomprising in addition silica, alkali metal and/or alkaline earth metalcomponents, provide improved performance and stability of operation.

The water absorption of the carrier may suitably be at least 0.2 g/g,preferably at least 0.25 g/g, more preferably at least 0.3 g/g, mostpreferably at least 0.35 g/g; and the water absorption may suitably beat most 0.85 g/g, preferably at most 0.7 g/g, more preferably at most0.65 g/g, most preferably at most 0.6 g/g. The water absorption of thecarrier may be in the range of from 0.2 to 0.85 g/g, preferably in therange of from 0.25 to 0.7 g/g, more preferably from 0.3 to 0.65 g/g,most preferably from 0.42 to 0.52 g/g. A higher water absorption may bein favor in view of a more efficient deposition of the metal andpromoters on the carrier by impregnation. However, at a higher waterabsorption, the carrier, or the catalyst made therefrom, may have lowercrush strength. As used herein, water absorption is deemed to have beenmeasured in accordance with ASTM C20, and water absorption is expressedas the weight of the water that can be absorbed into the pores of thecarrier, relative to the weight of the carrier.

A carrier may be washed, to remove soluble residues, before depositionof the catalyst ingredients on the carrier. Additionally, the materialsused to form the carrier, including the burnout materials, may be washedto remove soluble residues. Such carriers are described in U.S. Pat. No.6,368,998 and W0-A2-2007/095453, which are incorporated herein byreference. On the other hand, unwashed carriers may also be usedsuccessfully. Washing of the carrier generally occurs under conditionseffective to remove most of the soluble and/or ionizable materials fromthe carrier.

The washing liquid may be, for example water, aqueous solutionscomprising one or more salts, or aqueous organic diluents. Suitablesalts for inclusion in an aqueous solution may include, for exampleammonium salts. Suitable ammonium salts may include, for exampleammonium nitrate, ammonium oxalate, ammonium fluoride, and ammoniumcarboxylates, such as ammonium acetate, ammonium citrate, ammoniumhydrogencitrate, ammonium formate, ammonium lactate, and ammoniumtartrate. Suitable salts may also include other types of nitrates suchas alkali metal nitrates, for example lithium nitrate, potassium nitrateand cesium nitrate. Suitable quantities of total salt present in theaqueous solution may be at least 0.001% w, in particular at least 0.005%w, more in particular at least 0.01% w and at most 10% w, in particularat most 1% w, for example 0.03% w. Suitable organic diluents which mayor may not be included are, for example, one or more of methanol,ethanol, propanol, isopropanol, tetrahydrofuran, ethylene glycol,ethylene glycol dimethyl ether, diethylene glycol dimethyl ether,dimethylformamide, acetone, or methyl ethyl ketone.

The preparation of the silver catalyst is known in the art and the knownmethods are applicable to the preparation of the catalyst which may beused in the practice of the present invention. Methods of depositingsilver on the carrier include impregnating the carrier or carrier bodieswith a silver compound containing cationic silver and/or complexedsilver and performing a reduction to form metallic silver particles. Forfurther description of such methods, reference may be made to U.S. Pat.Nos. 5,380,697, 5,739,075, 4,766,105, and U.S. Pat. No. 6,368,998, whichare incorporated herein by reference. Suitably, silver dispersions, forexample silver sols, may be used to deposit silver on the carrier.

The reduction of cationic silver to metallic silver may be accomplishedduring a step in which the catalyst is dried, so that the reduction assuch does not require a separate process step. This may be the case ifthe silver containing impregnation solution comprises a reducing agent,for example, an oxalate, a lactate or formaldehyde.

Appreciable catalytic activity is obtained by employing a silver contentof the catalyst of at least 10 g/kg, relative to the weight of thecatalyst. Preferably, the catalyst comprises silver in a quantity offrom 10 to 500 g/kg, more preferably from 50 to 450 g/kg, for example105 g/kg, or 120 g/kg, or 190 g/kg, or 250 g/kg, or 350 g/kg. As usedherein, unless otherwise specified, the weight of the catalyst is deemedto be the total weight of the catalyst including the weight of thecarrier and catalytic components.

In an embodiment, the catalyst employs a silver content of the catalystof at least 150 g/kg, relative to the weight of the catalyst.Preferably, the catalyst comprises silver in a quantity of from 150 to500 g/kg, more preferably from 170 to 450 g/kg, for example 190 g/kg, or250 g/kg, or 350 g/kg.

The catalyst for use in the present invention additionally comprises arhenium promoter component. The form in which the rhenium promoter maybe deposited onto the carrier is not material to the invention. Forexample, the rhenium promoter may suitably be provided as an oxide or asan oxyanion, for example, as a rhenate or perrhenate, in salt or acidform.

The rhenium promoter may be present in a quantity of at least 0.01mmole/kg, preferably at least 0.1 mmole/kg, more preferably at least 0.5mmole/kg, most preferably at least 1 mmole/kg, in particular at least1.25 mmole/kg, more in particular at least 1.5 mmole/kg, calculated asthe total quantity of the element relative to the weight of thecatalyst. The rhenium promoter may be present in a quantity of at most500 mmole/kg, preferably at most 50 mmole/kg, more preferably at most 10mmole/kg, calculated as the total quantity of the element relative tothe weight of the catalyst.

In an embodiment, the rhenium promoter is present in a quantity of atleast 1.75 mmole/kg, preferably at least 2 mmole/kg, calculated as thetotal quantity of the element relative to the weight of the catalyst.The rhenium promoter may be present in a quantity of at most 15mmole/kg, preferably at most 10 mmole/kg, more preferably at most 8mmole/kg, calculated as the total quantity of the element relative tothe weight of the catalyst.

In an embodiment, the catalyst may further comprise a potassium promoterdeposited on the carrier. The potassium promoter may be deposited in aquantity of at least 0.5 mmole/kg, preferably at least 1 mmole/kg, morepreferably at least 1.5 mmole/kg, most preferably at least 1.75mmole/kg, calculated as the total quantity of the potassium elementdeposited relative to the weight of the catalyst. The potassium promotermay be deposited in a quantity of at most 20 mmole/kg, preferably atmost 15 mmole/kg, more preferably at most 10 mmole/kg, most preferablyat most 5 mmole/kg, on the same basis. The potassium promoter may bedeposited in a quantity in the range of from 0.5 to 20 mmole/kg,preferably from 1 to 15 mmole/kg, more preferably from 1.5 to 7.5mmole/kg, most preferably from 1.75 to 5 mmole/kg, on the same basis. Acatalyst prepared in accordance with the present invention can exhibitan improvement in selectivity, activity, and/or stability of thecatalyst especially when operated under conditions where the reactionfeed contains low levels of carbon dioxide.

The catalyst for use in the present invention may additionally comprisea rhenium co-promoter. The rhenium co-promoter may be selected fromtungsten, molybdenum, chromium, sulfur, phosphorus, boron, and mixturesthereof.

The rhenium co-promoter may be present in a total quantity of at least0.1 mmole/kg, more typically at least 0.25 mmole/kg, and preferably atleast 0.5 mmole/kg, calculated as the element (i.e. the total oftungsten, chromium, molybdenum, sulfur, phosphorus and/or boron),relative to the weight of the catalyst. The rhenium co-promoter may bepresent in a total quantity of at most 40 mmole/kg, preferably at most10 mmole/kg, more preferably at most 5 mmole/kg, on the same basis. Theform in which the rhenium co-promoter may be deposited on the carrier isnot material to the invention. For example, it may suitably be providedas an oxide or as an oxyanion, for example, as a sulfate, borate ormolybdate, in salt or acid form.

In an embodiment, the catalyst contains the rhenium promoter andtungsten in a molar ratio of the rhenium promoter to tungsten of greaterthan 2, more preferably at least 2.5, most preferably at least 3. Themolar ratio of the rhenium promoter to tungsten may be at most 20,preferably at most 15, more preferably at most 10.

In an embodiment, the catalyst comprises the rhenium promoter andadditionally a first co-promoter component and a second co-promotercomponent. The first co-promoter may be selected from sulfur,phosphorus, boron, and mixtures thereof. It is particularly preferredthat the first co-promoter comprises, as an element, sulfur. The secondco-promoter component may be selected from tungsten, molybdenum,chromium, and mixtures thereof. It is particularly preferred that thesecond co-promoter component comprises, as an element, tungsten and/ormolybdenum, in particular tungsten. The form in which the firstco-promoter and second co-promoter components may be deposited onto thecarrier is not material to the invention. For example, the firstco-promoter and second co-promoter components may suitably be providedas an oxide or as an oxyanion, for example, as a tungstate, molybdate,or sulfate, in salt or acid form.

In this embodiment, the first co-promoter may be present in a totalquantity of at least 0.2 mmole/kg, preferably at least 0.3 mmole/kg,more preferably at least 0.5 mmole/kg, most preferably at least 1mmole/kg, in particular at least 1.5 mmole/kg, more in particular atleast 2 mmole/kg, calculated as the total quantity of the element (i.e.,the total of sulfur, phosphorus, and/or boron) relative to the weight ofthe catalyst. The first co-promoter may be present in a total quantityof at most 50 mmole/kg, preferably at most 40 mmole/kg, more preferablyat most 30 mmole/kg, most preferably at most 20 mmole/kg, in particularat most 10 mmole/kg, more in particular at most 6 mmole/kg, calculatedas the total quantity of the element relative to the weight of thecatalyst.

In this embodiment, the second co-promoter component may be present in atotal quantity of at least 0.1 mmole/kg, preferably at least 0.15mmole/kg, more preferably at least 0.2 mmole/kg, most preferably atleast 0.25 mmole/kg, in particular at least 0.3 mmole/kg, more inparticular at least 0.4 mmole/kg, calculated as the total quantity ofthe element (i.e., the total of tungsten, molybdenum, and/or chromium)relative to the weight of the catalyst. The second co-promoter may bepresent in a total quantity of at most 40 mmole/kg, preferably at most20 mmole/kg, more preferably at most 10 mmole/kg, most preferably atmost 5 mmole/kg, calculated as the total quantity of the elementrelative to the weight of the catalyst.

In an embodiment, the molar ratio of the first co-promoter to the secondco-promoter may be greater than 1. In this embodiment, the molar ratioof the first co-promoter to the second co-promoter may preferably be atleast 1.25, more preferably at least 1.5, most preferably at least 2, inparticular at least 2.5. The molar ratio of the first co-promoter to thesecond co-promoter may be at most 20, preferably at most 15, morepreferably at most 10.

In an embodiment, the molar ratio of the rhenium promoter to the secondco-promoter may be greater than 1. In this embodiment, the molar ratioof the rhenium promoter to the second co-promoter may preferably be atleast 1.25, more preferably at least 1.5. The molar ratio of the rheniumpromoter to the second co-promoter may be at most 20, preferably at most15, more preferably at most 10.

In an embodiment, the catalyst comprises the rhenium promoter in aquantity of greater than 1 mmole/kg, relative to the weight of thecatalyst, and the total quantity of the first co-promoter and the secondco-promoter deposited on the carrier may be at most 12 mmole/kg,calculated as the total quantity of the elements (i.e., the total ofsulfur, phosphorous, boron, tungsten, molybdenum and/or chromium)relative to the weight of the catalyst. In this embodiment, the totalquantity of the first co-promoter and the second co-promoter maypreferably be at most 10 mmole/kg, more preferably at most 8 mmole/kg ofcatalyst. In this embodiment, the total quantity of the firstco-promoter and the second co-promoter may preferably be at least 0.1mmole/kg, more preferably at least 0.5 mmole/kg, most preferably atleast 1 mmole/kg of the catalyst.

The catalyst may preferably further comprise a further element depositedon the carrier. Eligible further elements may be one or more ofnitrogen, fluorine, alkali metals, alkaline earth metals, titanium,hafnium, zirconium, vanadium, thallium, thorium, tantalum, niobium,gallium and germanium and mixtures thereof. Preferably, the alkalimetals are selected from lithium, sodium and/or cesium. Preferably, thealkaline earth metals are selected from calcium, magnesium and barium.Preferably, the further element may be present in the catalyst in atotal quantity of from 0.01 to 500 mmole/kg, more preferably from 0.5 to100 mmole/kg, calculated as the total quantity of the element relativeto the weight of the catalyst. The further element may be provided inany form. For example, salts or hydroxides of an alkali metal or analkaline earth metal are suitable. For example, lithium compounds may belithium hydroxide or lithium nitrate.

In an embodiment, the catalyst may comprise cesium as a further elementin a quantity of more than 1.0 mmole/kg, in particular at least 2.0mmole/kg, more in particular at least 3.0 mmole/kg, calculated as thetotal quantity of the element relative to the weight of the catalyst. Inthis embodiment, the catalyst may comprise cesium in a quantity of atmost 20 mmole/kg, in particular at most 15 mmole/kg, calculated as thetotal quantity of the element relative to the weight of the catalyst Asused herein, unless otherwise specified, the quantity of alkali metalpresent in the catalyst and the quantity of water leachable componentspresent in the carrier are deemed to be the quantity insofar as it canbe extracted from the catalyst or carrier with de-ionized water at 100°C. The extraction method involves extracting a 10-gram sample of thecatalyst or carrier three times by heating it in 20 ml portions ofde-ionized water for 5 minutes at 100° C. and determining in thecombined extracts the relevant metals by using a known method, forexample atomic absorption spectroscopy.

As used herein, unless otherwise specified, the quantity of alkalineearth metal present in the catalyst and the quantity of acid leachablecomponents present in the carrier are deemed to be the quantity insofaras it can be extracted from the catalyst or carrier with 10% w nitricacid in de-ionized water at 100° C. The extraction method involvesextracting a 10-gram sample of the catalyst or carrier by boiling itwith a 100 ml portion of 10% w nitric acid for 30 minutes (1 atm., i.e.101.3 kPa) and determining in the combined extracts the relevant metalsby using a known method, for example atomic absorption spectroscopy.Reference is made to U.S. Pat. No. 5,801,259, which is incorporatedherein by reference.

Ethylene oxide produced may be recovered from the product mix by usingmethods known in the art, for example by absorbing ethylene oxide from areactor outlet stream in water and optionally recovering ethylene oxidefrom the aqueous solution by distillation. At least a portion of theaqueous solution containing ethylene oxide may be applied in asubsequent process for converting ethylene oxide into a 1,2-diol, a1,2-diol ether, a 1,2-carbonate, or an alkanolamine, in particularethylene glycol, ethylene glycol ethers, ethylene carbonate, or alkanolamines.

Ethylene oxide produced in the epoxidation process may be converted intoa 1,2-diol, a 1,2-diol ether, a 1,2-carbonate, or an alkanolamine. Asthis invention leads to a more attractive process for the production ofethylene oxide, it concurrently leads to a more attractive process whichcomprises producing ethylene oxide in accordance with the invention andthe subsequent use of the obtained ethylene oxide in the manufacture ofthe 1,2-diol, 1,2-diol ether, 1,2-carbonate, and/or alkanolamine.

The conversion into the 1,2-diol (i.e., ethylene glycol) or the 1,2-diolether (i.e., ethylene glycol ethers) may comprise, for example, reactingethylene oxide with water, suitably using an acidic or a basic catalyst.For example, for making predominantly the 1,2-diol and less 1,2-diolether, ethylene oxide may be reacted with a ten fold molar excess ofwater, in a liquid phase reaction in presence of an acid catalyst, e.g.0.5-1.0% w sulfuric acid, based on the total reaction mixture, at 50-70°C. at 1 bar absolute, or in a gas phase reaction at 130-240° C. and20-40 bar absolute, preferably in the absence of a catalyst. Thepresence of such a large quantity of water may favor the selectiveformation of 1,2-diol and may function as a sink for the reactionexotherm, helping control the reaction temperature. If the proportion ofwater is lowered, the proportion of 1,2-diol ethers in the reactionmixture is increased. Alternative 1,2-diol ethers may be prepared byconverting ethylene oxide with an alcohol, in particular a primaryalcohol, such as methanol or ethanol, by replacing at least a portion ofthe water by the alcohol.

Ethylene oxide may be converted into the corresponding 1,2-carbonate byreacting ethylene oxide with carbon dioxide. If desired, ethylene glycolmay be prepared by subsequently reacting the 1,2-carbonate with water oran alcohol to form the glycol. For applicable methods, reference is madeto U.S. Pat. No. 6,080,897, which is incorporated herein by reference.

The conversion into the alkanolamine may comprise, for example, reactingethylene oxide with ammonia. Anhydrous ammonia is typically used tofavor the production of monoalkanolamine. For methods applicable in theconversion of ethylene oxide into the alkanolamine, reference may bemade to, for example U.S. Pat. No. 4,845,296, which is incorporatedherein by reference.

The 1,2-diol and the 1,2-diol ether may be used in a large variety ofindustrial applications, for example in the fields of food, beverages,tobacco, cosmetics, thermoplastic polymers, curable resin systems,detergents, heat transfer systems, etc. The 1,2-carbonates may be usedas a diluent, in particular as a solvent. The alkanolamine may be used,for example, in the treating (“sweetening”) of natural gas.

EXAMPLE 1

A “tube fraction” is a sample of catalyst that has been removed from acommercial EO reactor after an extended period of commercial operation,with each fraction representing a specific axial region of a specificreactor tube. In Example 1, an assortment of tube fractions weresubjected to microreactor testing at the conclusion of a commercial highselectivity catalyst cycle at Plant A. The following parameters wereused in our tests in order to closely simulate the average commercialoperation conditions for Plant A: Feedstock: 27.5% v C₂H₄, 7.3% v O₂,1.4% v CO₂; variable amounts of ethyl chloride moderator, balance N₂.Conditions: 4680 hr⁻¹ Gas Hourly Space Velocity, 19.5 barg pressure, 253kg/m³/hr workrate.

To demonstrate the instant invention in our microreactors, a two-steptesting protocol was utilized. It is important to appreciate that Step 1is not the inventive step per se, but rather, provides criticalinformation that is required prior to performing the inventive step. Theobjective of Step 1 was to identify the “chloride optimum” for a givencatalyst as it operates at the above-specified conditions in ourmicroreactor, which is to say, to identify the concentration of ethylchloride “EC-opt” for which the maximum selectivity is achieved at thespecified conditions. A properly operated commercial plant shouldalready know the optimum level of chloride for their catalyst at theircurrent conditions. However, the chloride optimum reported by the plantfrequently does not exactly correspond to the chloride optimum that weobserve in our microreactor testing. For this reason, unlike acommercial plant wishing to employ the instant invention, Step 1 must beemployed to identify the chloride optimum of each sample in ourmicroreactors prior to performing the instant invention. To employ theinstant invention, a commercial plant needs only apply the proceduredescribed below as Step 2.

To properly identify the chloride optimum, one must take care to beginthe optimization procedure at a chloride level that is below the optimumlevel and continue stepwise increasing the chloride level until abovethe optimum level, collecting equilibrated performance data at eachchloride level, and thereby tracing out a selectivity-versus-chloridecurve that displays a region of maximum selectivity lying somewherebetween the terminal points of the curve.

A catalyst is said to be “undermoderated” when the amount of chloride onthe catalyst surface and the amount of chloride in the gas phase are inequilibrium, and said chloride level on the catalyst surface is lessthan the amount required to achieve maximum selectivity. To render thecatalyst into an undermoderated state, one can select from manycombinations of catalyst temperature and chloride level in the gasphase. For example, one can hold the catalyst temperature constant atits “normal” operating level, and reduce or entirely eliminate thechloride content of the gas stream for a period of time. In thatscenario, lower levels of chloride are preferable, since the timerequired to drive the catalyst into an undermoderated state is relatedto the degree of undermoderation. As a second example, one can hold thechloride concentration of the feed stream at its “normal” operatinglevel, and elevate the temperature of the catalyst. As a third example,one can manipulate both the temperature and the chloride level in orderto induce an undermoderated state.

In each test comprising Example 1, operation commenced at a constantcatalyst temperature of 240° C. while including 1.2 ppmv moderator inthe feed stream for 8 hours. Moderator flow was reduced to zero whilecatalyst temperature was held at 240° C. for a period of 4 hours, andthen sequentially increased to 247° C. for a period of 4 hours, and to255° C. for a period of 8 hours. Moderator flow was then re-introducedat a level of 1.5 ppmv for a period of 36 hours, at which time thereactor was placed on workrate control at the target workrate. Themoderator level was then sequentially incremented every 36 hours to 1.8,2.1, 2.4, 2.7 and 3.0 ppmv in order to identify the optimalconcentration of chloride “EC-opt”, the corresponding or “optimized”selectivity, and the corresponding catalyst temperature. In the Table,optimized catalyst performance data are shown in the columns labeled“Performance Prior to Temperature Treatment”.

Step 2 of the testing protocol was to perform the steps comprising theinstant invention. The catalyst was again subjected to a chloride stripat 240° C. for a period of 4 hours, 247° C. for a period of 4 hours, andthen 255° C. for a period of 8 hours. Following the chloride strip,moderator flow was re-introduced. However, unlike the protocol formoderator introduction protocol described earlier, this time themoderator flow was re-introduced precisely at the level that hadpreviously been identified as “optimal”. After selectivity andtemperature reached equilibrium, performance data were againrecorded—see columns labeled “Performance Following TemperatureTreatment”. The improvement that we observed following Step 2 is shownin the column “Improvement in Selectivity”.

EXAMPLE 1 - Commercially Aged High Selectivity Catalyst. PerformancePerformance Prior To Following Micro- Temperature Temperature reactorTreatment Treatment Improve- Bed No. Selec- Catalyst Selec- Catalystment In Catalyst LR-27126 tivity Temp tivity Temp Selectivity SampleLR*27139 (%) (° C.) (%) (° C.) (%) 1 -104-2 81.9 259 84.1 256 2.2 2-104-3 81.1 259 83.2 259 2.1 3 -104-1 81.7 243 84.1 246 2.4 4 -104-481.7 248 83.1 251 1.4 5 -133-2 81.5 255 82.4 256 0.9 6 *003-2 81.1 25783.2 254 2.1 7 *003-3 81.2 253 83.6 251 2.4 8 *003-1 79.8 244 82.2 2452.4 9 *003-4 80.6 247 83.1 247 2.5 10 *007-1 81.2 250 83.8 249 2.6 11*007-2 80.4 250 82.3 249 1.9 12 *007-3 81.6 250 82.3 248 0.7 Average allsections: 81.2 251 83.1 251 2.0

EXAMPLE 2

In Example 2, the same protocol as outlined above was followed for adifferent set of commercially aged high selectivity catalyst samplesfrom Plant B. Results are summarized in the Table “Example2—Commercially Aged High Selectivity Catalyst”.

EXAMPLE 2 - Commercially Aged High Selectivity Catalyst. PerformancePerformance Prior To Following Micro- Temperature Temperature reactorTreatment Treatment Improve- Bed No. Selec- Catalyst Selec- Catalystment In Catalyst LR-27126 tivity Temp tivity Temp Selectivity SampleLR*27139 (%) (° C.) (%) (° C.) (%) 1 -014-2 79.7 244 82.5 242 2.8 2*138-4 79.5 242 82.0 241 2.5 3 -036-3 80.5 243 83.4 240 2.9 4 -026-181.1 242 83.6 242 2.5 5 -026-4 81.8 242 84.0 240 2.2 Average allsections: 80.5 242 83.1 241 2.6

EXAMPLE 3

In Example 3 an experiment was run to show the effect of increasing thetemperature to affect a hard strip. In this experiment 4.6 grams of arhenium containing high selectivity catalyst according to US2009/0281345A1 was loaded into a U-tube reactor at 14-20 mesh size. Thecatalyst was operated for 10 months at 30 mole-% ethylene, 8 mole-%oxygen, 1 mole-% carbon dioxide with the balance nitrogen and a reactorinlet pressure of 210 psig, with an EO concentration of 3.09%. The ethylchloride (EC) moderator was adjusted periodically to maintain maximumselectivity. After 10 months of operation the selectivity declined to88%, the EC concentration was 4.2 ppm and the catalyst temperature was249.2° C. A hard strip was performed according to the claimed inventionby increasing the temperature to 250.2° C. and decreasing the ECconcentration to ZERO for 6 hours and then resetting the ECconcentration to 3.9 ppm. The catalyst temperature was adjusted to make3.09% EO. The resulting catalyst selectivity was 89.4% at 249.2° C., anINCREASE of 1.4% with no change in temperature.

What is claimed is:
 1. A process for improving the selectivity of anethylene epoxidation process employed in a reactor, said processcomprising: (a) contacting an epoxidation catalyst comprising silver andrhenium with a feed comprising ethylene, oxygen, and an organic chloridefor a period of time T₁ sufficient to produce at least 0.1 kilotons ofethylene oxide per cubic meter of epoxidation catalyst; (b) subsequentto step (a) and over a period of time T₂, (i) reducing the quantity oforganic chloride added to the feed by at least 50% and (ii) increasingthe reactor temperature by about 1 to 30° C.; and (c) subsequent to step(b), increasing the quantity of organic chloride added to the feed suchthat the quantity present in the feed is between 80 and 120% of thequantity that was present in the feed prior to step (b), and reducingthe reactor temperature to within 5° C. of the prior reactortemperature.
 2. The process of claim 1 wherein the organic chloride isselected from the group consisting of methyl chloride, ethyl chloride,ethylene dichloride, vinyl chloride and mixtures thereof.
 3. The processof claim 2 wherein the quantity of organic chloride present in the feedin step (a) is in the range of from 0.01×10⁻⁴ to 2×10⁻³ mole-%,calculated as moles of chloride, relative to the total feed.
 4. Theprocess of claim 3 wherein in step (b), the quantity of organic chlorideadded to the feed is reduced by at least 75%.
 5. The process of claim 4wherein in step (b), the reactor temperature is increased by about 2 to15° C.
 6. The process of claim 1 wherein T₂ is about 2 hours to about 72hours.
 7. The process of claim 1 wherein T₂ is about 4 to 24 hours. 8.The process of claim 1 wherein rhenium is present in the epoxidationcatalyst in an amount greater than 1 mmole/kg, relative to the weight ofthe catalyst, and the epoxidation catalyst further comprises (a) a firstco-promoter selected from the group consisting of sulfur, phosphorus,boron, and mixtures thereof; and (b) a second co-promoter selected fromthe group consisting of tungsten, molybdenum, chromium, and mixturesthereof.
 9. The process of claim 8 wherein the total quantity of thefirst co-promoter and the second co-promoter is at most 10.0 mmole/kg,relative to the weight of the catalyst; and wherein the epoxidationcatalyst further comprises a carrier having a monomodal, bimodal ormultimodal pore size distribution, with a pore diameter range of0.01-200 μm, a specific surface area of 0.03-10 m²/g, a pore volume of0.2-0.7 cm³/g, wherein the median pore diameter of said carrier is0.1-100 μm and has a water absorption of 10-80%.
 10. The process ofclaim 1 wherein the feed further comprises CO₂ in a quantity that isless than 2%.
 11. The process of claim 1 further comprising reacting atleast a portion the ethylene oxide with at least one selected from thegroup consisting of water, an alcohol, carbon dioxide and an amine toform, respectively, a 1,2-diol, a 1,2-diol ether, a 1,2-carbonate, or analkanolamine.
 12. A process comprising: (a) contacting an epoxidationcatalyst comprising silver and rhenium with a feed comprising ethylene,oxygen, and an organic chloride for a period of time sufficient toproduce at least 0.1 kilotons of ethylene oxide per cubic meter ofepoxidation catalyst; (b) subsequent to step (a) and over a period ofabout 2 to 72 hours, (i) reducing the quantity of organic chloride addedto the feed by at least 50% and (ii) increasing the reactor temperatureby about 1 to 30° C.; and (c) subsequent to step (b), increasing thequantity of organic chloride added to the feed such that the quantitypresent in the feed is between 80 and 120% of the quantity that waspresent in the feed prior to step (b), and reducing the reactortemperature to within 5° C. of the prior reactor temperature.